U.S. patent number 9,708,479 [Application Number 14/901,948] was granted by the patent office on 2017-07-18 for multistage process for producing polyethylene compositions.
This patent grant is currently assigned to BOREALIS AG. The grantee listed for this patent is BOREALIS AG. Invention is credited to Alexandra Romina Albunia, Thomas Hjertberg, Dusan Jeremic, Yi Liu, Tanja Piel, Floran Prades, Joseph Thorman, Ravindra Tupe, Udo Wahner.
United States Patent |
9,708,479 |
Hjertberg , et al. |
July 18, 2017 |
Multistage process for producing polyethylene compositions
Abstract
A process for polymerizing ethylene in the presence of a
polymerization catalyst by copolymerizing ethylene with a comonomer
selected from the group of C4-C10 alpha-olefins in three
polymerization stages. The polymers produced in the three stages
have different melt flow rates. The polymer composition produced by
the process has good mechanical properties and can be used for
making pipes. The process has a good productivity and provides a
stable and economic operation.
Inventors: |
Hjertberg; Thomas (Kungshamn,
SE), Tupe; Ravindra (Maharashtra, IN),
Wahner; Udo (Linz, AT), Piel; Tanja (Linz,
AT), Liu; Yi (Engerwitzdorf, AT), Albunia;
Alexandra Romina (Linz, AT), Prades; Floran
(Linz, AT), Thorman; Joseph (Helsinki, FI),
Jeremic; Dusan (Linz, AT) |
Applicant: |
Name |
City |
State |
Country |
Type |
BOREALIS AG |
Vienna |
N/A |
AT |
|
|
Assignee: |
BOREALIS AG (Vienna,
AT)
|
Family
ID: |
49328307 |
Appl.
No.: |
14/901,948 |
Filed: |
September 24, 2014 |
PCT
Filed: |
September 24, 2014 |
PCT No.: |
PCT/EP2014/002587 |
371(c)(1),(2),(4) Date: |
December 29, 2015 |
PCT
Pub. No.: |
WO2015/051878 |
PCT
Pub. Date: |
April 16, 2015 |
Prior Publication Data
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|
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Document
Identifier |
Publication Date |
|
US 20160369091 A1 |
Dec 22, 2016 |
|
Foreign Application Priority Data
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|
|
|
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Oct 10, 2013 [EP] |
|
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13004879 |
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Current U.S.
Class: |
1/1 |
Current CPC
Class: |
C08F
2/001 (20130101); C08L 23/0815 (20130101); C08L
23/06 (20130101); C08F 210/16 (20130101); C08F
2/001 (20130101); C08L 23/0815 (20130101); C08L
23/06 (20130101); C08L 23/06 (20130101); C08L
23/06 (20130101); C08L 23/0815 (20130101); C08L
23/06 (20130101); C08L 2205/03 (20130101); C08L
2205/025 (20130101); C08L 2205/02 (20130101); C08F
210/16 (20130101); C08F 210/14 (20130101); C08F
2500/07 (20130101); C08F 2500/12 (20130101) |
Current International
Class: |
C08L
23/08 (20060101); C08F 2/00 (20060101) |
References Cited
[Referenced By]
U.S. Patent Documents
Foreign Patent Documents
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684871 |
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1415999 |
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1591460 |
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9618662 |
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WO |
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WO |
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2007025640 |
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Mar 2007 |
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WO |
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2013/144324 |
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Oct 2013 |
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WO |
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Other References
Geldart and Bayens: The Design of Distributors for Gas-fluidized
Beds, Powder Technology, vol. 42, 1985. cited by applicant .
Geldart: Gas Fluidization Technology, J. Wiley & Sons, 1986, p.
17-18, 169-186, 183. cited by applicant .
International Search Report and Written Opinion of Application No.
PCT/EP2014/002587 dated Oct. 27, 2015. cited by applicant.
|
Primary Examiner: Nutter; Nathan M
Attorney, Agent or Firm: Roberts Mlotkowski Safran Cole
& Calderon, P.C.
Claims
The invention claimed is:
1. A process for producing multimodal ethylene polymers in the
presence of a polymerization catalyst comprising the steps of:
polymerizing ethylene in a first polymerization stage at a
temperature of from 30 to 70.degree. C. and a pressure of from 1 to
150 bar in the presence of the polymerization catalyst to produce a
first homo- or copolymer of ethylene; polymerizing ethylene in a
second polymerization stage at a temperature of from 50 to
115.degree. C., and a pressure of from 1 to 150 bar in the presence
of hydrogen and the first homo- or copolymer of ethylene to produce
a first ethylene polymer mixture comprising the first homo- or
copolymer of ethylene and a second homopolymer of ethylene;
copolymerizing ethylene in a third polymerization stage at a
temperature of from 50 to 100.degree. C., and a pressure of from 10
to 40 bar in the presence of an alpha-olefin comonomer having from
4 to 10 carbon atoms and the first ethylene polymer mixture to
produce a second ethylene polymer mixture comprising the first
ethylene polymer mixture and a third copolymer of ethylene; wherein
the first homo- or copolymer of ethylene has a melt flow rate
MFR.sub.21 of not more than 1.0 g/10 min; the first ethylene
polymer mixture has a melt flow rate MFR.sub.2 of from 10 to 250
g/10 min; and the second ethylene polymer mixture has a melt flow
rate MFR.sub.21 of from 1 to 50 g/10 min and a density of from 935
to 960 kg/m.sup.3.
2. The process according to claim 1 wherein the second ethylene
polymer mixture comprises from 40 to 60% by weight of the first
ethylene polymer mixture and from 40 to 60% by weight of the third
copolymer of ethylene.
3. The process according to claim 1 wherein the second ethylene
polymer mixture comprises from 45 to 60% by weight of the first
ethylene polymer mixture and from 40 to 55% by weight of the third
copolymer of ethylene.
4. The process according to claim 1 wherein the first ethylene
polymer mixture comprises from 2 to 20% by weight of the first
homo- or copolymer of ethylene and 80 to 98% by weight of the
second copolymer of ethylene.
5. The process according to claim 1, wherein the ethylene content
in the fluid phase of the first reactor is 0.2 to 50% by mole.
6. The process according to claim 1, wherein the density of the
first polymer mixture is higher than 965 kg/m.sup.3.
7. The process according to claim 1, wherein the density of the
second polymer mixture 935 to 960 kg/m.sup.3.
8. The process according to claim 1, wherein the MFR.sub.21 of the
first homo- or copolymer of ethylene is from 0.01 to 1.0 g/10
min.
9. The process according to claim 1, wherein the MFR.sub.2 of the
first polymer mixture is from 2 to 200 g/10 min.
10. The process according to claim 1, wherein the MFR.sub.2 of the
first polymer mixture is higher than the MFR.sub.2 of the second
polymer mixture.
11. The process according to claim 1, wherein the first and second
polymerization stages are conducted as a slurry polymerization.
12. The process according to claim 1, wherein the third
polymerization stage is conducted in vapour phase in a fluidized
bed.
13. The process according to claim 1, wherein the process comprises
the steps of: continuously introducing ethylene, diluent and the
polymerization catalyst into the first polymerization stage;
continuously withdrawing a stream comprising the first homo- or
copolymer of ethylene from the first polymerization stage and
passing it into the second polymerization stage; continuously
introducing streams of ethylene, hydrogen and diluent into the
second polymerization stage; continuously withdrawing a stream
comprising the first ethylene polymer mixture from the second
polymerization stage and passing at least part of the stream to the
third polymerization stage; continuously introducing streams of
ethylene, hydrogen and the alpha-olefin comonomer into the third
polymerization stage; continuously withdrawing a stream comprising
the second ethylene polymer mixture from the third polymerization
stage and recovering the second ethylene polymer mixture
therefrom.
14. The process according to claim 1, wherein the polymerization
catalyst comprises (i) a solid catalyst component comprising a
magnesium compound and a halogen containing titanium compound
supported on an inorganic oxide support and (ii) an activator.
15. The process according to claim 14, wherein the inorganic oxide
support has a median particle size (D50) of from 5 to 15 .mu.m.
Description
FIELD OF THE INVENTION
The present invention is directed to a method of producing ethylene
polymers. Especially, the present invention is directed to a method
of making multimodal ethylene polymers where the method comprises
polymerizing ethylene in three polymerization stages.
Problem to be Solved
It is known in the art to polymerize ethylene in three stages from
WO-A-1996018662 where ethylene is first (co)polymerized in a
prepolymerization step in a loop reactor and subsequently
polymerized in loop and gas phase reactors. In the first step a
polymer having an MFR.sub.2 between 0.01 and 50 g/10 min and a
density between 925 and 975 kg/m.sup.3 was produced. The amount of
the polymer produced in the first stage was from 1 to 20%,
preferably from 5 to 15% of the total production in all three
reactors.
EP-B-1228101 discloses a process where ethylene is first
homopolymerized to produce a low molecular weight component, then
copolymerized to produce a high molecular weight component and then
homo- or copolymerized to produce an ultra-high molecular weight
component. The relative amounts of the low molecular weight
component, high molecular weight component and the ultra-high
molecular weight component were from 30 to 60%, from 30 to 65% and
from 1 to 30%, respectively.
In spite of the methods disclosed in the prior art, there still
exists a need for a process being able to produce multimodal HDPE
compositions having a broad molecular weight distribution which
process is stable and reliable to operate and where the polymer
powder has good handling properties.
The objective of the present invention is thus to overcome the
problems of prior art processes and to provide a process which
allows the production of ethylene polymer compositions having a
broad molecular weight distribution with a high production rate,
high productivity of the catalyst and a good production economy.
The resulting polymer compositions have a combination of excellent
processability, good mechanical properties and good
homogeneity.
SUMMARY OF THE INVENTION
The present invention provides a process for producing multimodal
ethylene polymers in the presence of a polymerization catalyst
comprising the steps of: polymerizing ethylene in a first
polymerization stage at a temperature of from 30 to 70.degree. C.,
preferably from 40 to 60.degree. C. and a pressure of from 1 to 150
bar, preferably from 10 to 100 bar, in the presence of the
polymerization catalyst to produce a first homo- or copolymer of
ethylene; polymerizing ethylene in a second polymerization stage at
a temperature of from 50 to 115.degree. C., preferably from 90 to
105.degree. C., and a pressure of from 1 to 150 bar, preferably
from 50 to 100 bar in the presence of hydrogen and the first homo-
or copolymer of ethylene to produce a first ethylene polymer
mixture comprising the first homo- or copolymer of ethylene and a
second homopolymer of ethylene; copolymerizing ethylene in a third
polymerization stage at a temperature of from 50 to 100.degree. C.,
preferably from 65 to 90.degree. C., and a pressure of from 10 to
40 bar, preferably from 15 to 30 bar, in the presence of an
alpha-olefin comonomer having from 4 to 10 carbon atoms and the
first ethylene polymer mixture to produce a second ethylene polymer
mixture comprising the first ethylene polymer mixture and a third
copolymer of ethylene; wherein the first ethylene homopolymer has a
melt flow rate MFR.sub.21 of not more than 1.0 g/10 min; the first
ethylene polymer mixture has a melt flow rate MFR.sub.2 of from 10
to 250 g/10 min; and the second ethylene polymer mixture has a melt
flow rate MFR.sub.21 of from 1 to 50 g/10 min and a density of from
935 to 960 kg/m.sup.3.
DETAILED DESCRIPTION
Even though the present invention relates to a three-stage process
for producing ethylene polymer compositions, it should be
understood that the process may contain additional polymerization
stages to the three stages disclosed above. It may contain
additional polymerization stages, such as a prepolymerization
stage, as long as the polymer produced in such additional stages
does not substantially influence the properties of the polymer.
Furthermore, any one of the three polymerization stages disclosed
above may be conducted as two or more sub-stages, provided that the
polymer produced in each such sub-stage as well as their mixture
matches the description for the polymer for the respective
stage.
However, it is preferred to conduct each of the first, second and
third polymerization stage as a single polymerization stage in
order to prevent the process from becoming unnecessarily complex.
Therefore, in the most preferred embodiment the process consists of
three polymerization stages which may be preceded a
prepolymerization stage.
The present invention is directed to producing polymers of ethylene
comprising three ethylene polymer components where each of the
components is produced in a dedicated polymerization stage. One of
the components is an ethylene homopolymer and two are copolymers of
ethylene with at least one comonomer selected from the group
consisting alpha-olefins having from 4 to 10 carbon atoms.
Preferably the comonomer is selected from the group of
alpha-olefins having from 4 to 8 carbon atoms and their mixtures.
Especially preferably the comonomer is 1-butene or 1-hexene or
their mixture.
Catalyst
The polymerisation is conducted in the presence of an olefin
polymerisation catalyst. The catalyst may be any catalyst which is
capable of producing the desired ethylene polymer. Suitable
catalysts are, among others, Ziegler-Natta catalysts based on a
transition metal, such as titanium, zirconium and/or vanadium
catalysts. Especially Ziegler-Natta catalysts are useful as they
can produce polymers within a wide range of molecular weight with a
high productivity.
Suitable Ziegler-Natta catalysts preferably contain a magnesium
compound, an aluminium compound and a titanium compound supported
on a particulate support.
The particulate support can be an inorganic oxide support, such as
silica, alumina, titania, silica-alumina and silica-titania.
Preferably, the support is silica.
The average particle size (D50) of the silica support can be
typically from 10 to 100 .mu.m. However, it has turned out that
special advantages can be obtained if the support has an average
particle size (D50) from 5 to 20 .mu.m, preferably from 5 to 15
.mu.m.
The magnesium compound is a reaction product of a magnesium dialkyl
and an alcohol. The alcohol is a linear or branched aliphatic
monoalcohol. Preferably, the alcohol has from 6 to 16 carbon atoms.
Branched alcohols are especially preferred, and 2-ethyl-1-hexanol
is one example of the preferred alcohols. The magnesium dialkyl may
be any compound of magnesium bonding to two alkyl groups, which may
be the same or different. Butyl-octyl magnesium is one example of
the preferred magnesium dialkyls.
The aluminium compound is chlorine containing aluminium alkyl.
Especially preferred compounds are aluminium alkyl dichlorides and
aluminium alkyl sesquichlorides.
The titanium compound is a halogen containing titanium compound,
preferably chlorine containing titanium compound. Especially
preferred titanium compound is titanium tetrachloride.
The catalyst can be prepared by sequentially contacting the carrier
with the above mentioned compounds, as described in EP-A-688794 or
WO-A-99/51646. Alternatively, it can be prepared by first preparing
a solution from the components and then contacting the solution
with a carrier, as described in WO-A-01/55230.
Another group of suitable Ziegler-Natta catalysts contain a
titanium compound together with a magnesium halide compound acting
as a support. Thus, the catalyst contains a titanium compound on a
magnesium dihalide, like magnesium dichloride. Such catalysts are
disclosed, for instance, in WO-A-2005/118655 and EP-A-810235.
Still a further type of Ziegler-Natta catalysts are catalysts
prepared by a method, wherein an emulsion is formed, wherein the
active components form a dispersed, i.e. a discontinuous phase in
the emulsion of at least two liquid phases. The dispersed phase, in
the form of droplets, is solidified from the emulsion, wherein
catalyst in the form of solid particles is formed. The principles
of preparation of these types of catalysts are given in
WO-A-2003/106510 of Borealis.
The Ziegler-Natta catalyst is used together with an activator.
Suitable activators are metal alkyl compounds and especially
aluminium alkyl compounds. These compounds include alkyl aluminium
halides, such as ethylaluminium dichloride, diethylaluminium
chloride, ethylaluminium sesquichloride, dimethylaluminium chloride
and the like. They also include trialkylaluminium compounds, such
as trimethylaluminium, triethylaluminium, tri-isobutylaluminium,
trihexylaluminium and tri-n-octylaluminium. Furthermore they
include alkylaluminium oxy-compounds, such as methylaluminiumoxane
(MAO), hexaisobutylaluminiumoxane (HIBAO) and
tetraisobutylaluminiumoxane (TIBAO). Also other aluminium alkyl
compounds, such as isoprenylaluminium, may be used. Especially
preferred activators are trialkylaluminiums, of which
triethylaluminium, trimethylaluminium and tri-isobutylaluminium are
particularly used.
The amount in which the activator is used depends on the specific
catalyst and activator. Typically triethylaluminium is used in such
amount that the molar ratio of aluminium to the transition metal,
like Al/Ti, is from 1 to 1000, preferably from 3 to 100 and in
particular from about 5 to about 30 mol/mol.
Prepolymerization
The polymerization steps may be preceded by a prepolymerization
step. The purpose of the prepolymerization is to polymerize a small
amount of polymer onto the catalyst at a low temperature and/or a
low monomer concentration. By prepolymerization it is possible to
improve the performance of the catalyst in slurry and/or modify the
properties of the final polymer. The prepolymerization step is
conducted in slurry.
Thus, the prepolymerization step may be conducted in a loop
reactor. The prepolymerization is then preferably conducted in an
inert diluent, typically a hydrocarbon diluent such as methane,
ethane, propane, n-butane, isobutane, pentanes, hexanes, heptanes,
octanes etc., or their mixtures. Preferably the diluent is a
low-boiling hydrocarbon having from 1 to 4 carbon atoms or a
mixture of such hydrocarbons.
The temperature in the prepolymerization step is typically from 0
to 100.degree. C., preferably from 20 to 95.degree. C. and more
preferably from 40 to 70.degree. C.
The pressure is not critical and is typically from 1 to 150 bar,
preferably from 10 to 100 bar.
The amount of monomer is typically such that from about 0.1 to 1000
grams of monomer per one gram of solid catalyst component is
polymerized in the prepolymerization step. As the person skilled in
the art knows, the catalyst particles recovered from a continuous
prepolymerization reactor do not all contain the same amount of
prepolymer. Instead, each particle has its own characteristic
amount which depends on the residence time of that particle in the
prepolymerization reactor. As some particles remain in the reactor
for a relatively long time and some for a relatively short time,
then also the amount of prepolymer on different particles is
different and some individual particles may contain an amount of
prepolymer which is outside the above limits. However, the average
amount of prepolymer on the catalyst typically is within the limits
specified above.
The molecular weight of the prepolymer may be controlled by
hydrogen as it is known in the art. Further, antistatic additive
may be used to prevent the particles from adhering to each other or
the walls of the reactor, as disclosed in WO-A-96/19503 and
WO-A-96/32420.
The catalyst components are preferably all introduced to the
prepolymerization step when a prepolymerization step is present.
However, where the solid catalyst component and the cocatalyst can
be fed separately it is possible that only a part of the cocatalyst
is introduced into the prepolymerization stage and the remaining
part into subsequent polymerization stages. Also in such cases it
is necessary to introduce so much cocatalyst into the
prepolymerization stage that a sufficient polymerization reaction
is obtained therein.
First Polymerization Stage
In the first polymerization stage a first homo- or copolymer of
ethylene is produced. This is done by introducing a polymerization
catalyst, optionally through the prepolymerization stage as
disclosed above, into the first polymerization stage together with
ethylene monomer.
The first homo- or copolymer of ethylene preferably has a melt flow
rate MFR.sub.21 of from 0.01 to 1.0 g/10 min, more preferably from
0.1 to 1.0 g/10 min, and most preferably from 0.1 to 0.7 g/10 min.
Typically, an MFR.sub.21/MFR.sub.2 value within the range of 40 to
100, such as 70, can be used for a typical polymer produced in the
first polymerization stage
In an alternative preferred embodiment of the invention the first
homo- or copolymer of ethylene has such a high molecular weight
that the melt flow rate is immeasurably low. In such a case
practically no polymer comes out of the melt indexer during the
time of the measurement under 21.6 kg load. Then the melt flow rate
MFR.sub.21 is taken as 0 g/10 min.
The first polymerization stage is preferably conducted as a slurry
polymerization. The slurry polymerization usually takes place in an
inert diluent, typically a hydrocarbon diluent such as methane,
ethane, propane, n-butane, isobutane, pentanes, hexanes, heptanes,
octanes etc., or their mixtures. Preferably the diluent is a
low-boiling hydrocarbon having from 1 to 4 carbon atoms or a
mixture of such hydrocarbons. An especially preferred diluent is
propane, possibly containing minor amount of methane, ethane and/or
butane.
The ethylene content in the fluid phase of the slurry may be from
0.2 to about 50% by mole, preferably from about 0.5 to about 20% by
mole and in particular from about 0.5 to about 15% by mole. When
the polymerization is conducted as a continuous process in a loop
reactor a range of from 0.5 to 10, preferably from 0.5 to 6% by
mole has given good results. The benefit of having a high ethylene
concentration is that the productivity of the catalyst is increased
but the drawback is that more ethylene then needs to be recycled
than if the concentration was lower.
The temperature in the first polymerization stage is typically from
30 to 70.degree. C., preferably from 40 to 70.degree. C. The
pressure is from 1 to 150 bar, preferably from 10 to 100 bar.
The slurry polymerization may be conducted in any known reactor
used for slurry polymerization. Such reactors include a continuous
stirred tank reactor and a loop reactor. It is especially preferred
to conduct the polymerization in loop reactor. In such reactors the
slurry is circulated with a high velocity along a closed pipe by
using a circulation pump. Loop reactors are generally known in the
art and examples are given, for instance, in U.S. Pat. No.
4,582,816, U.S. Pat. No. 3,405,109, U.S. Pat. No. 3,324,093,
EP-A-479186 and U.S. Pat. No. 5,391,654.
The slurry may be withdrawn from the reactor either continuously or
intermittently. A preferred way of intermittent withdrawal is the
use of settling legs where slurry is allowed to concentrate before
withdrawing a batch of the concentrated slurry from the reactor.
The use of settling legs is disclosed, among others, in U.S. Pat.
No. 3,374,211, U.S. Pat. No. 3,242,150 and EP-A-1310295. Continuous
withdrawal is disclosed, among others, in EP-A-891990,
EP-A-1415999, EP-A-1591460 and WO-A-2007/025640. The continuous
withdrawal is advantageously combined with a suitable concentration
method, as disclosed in EP-A-1310295 and EP-A-1591460. It is
preferred to withdraw the slurry from the first polymerization
stage continuously.
Hydrogen is typically not introduced into the first polymerization
stage. However, in some cases there may be a need to feed a small
stream of hydrogen into the first polymerization stage to control
the molecular weight of the polymer as known in the art. If
hydrogen is used then the molar ratio of hydrogen to ethylene
should not exceed 10 mol/kmol (or mol/1000 mol) and preferably not
5 mol/kmol.
It is possible to feed a small amount of comonomer into the first
polymerization stage. The comonomer, if used, is an alpha-olefin
having from 4 to 10 carbon atoms. The molar ratio of the comonomer
to ethylene is not higher than 100 mol/kmol, preferably not higher
than 50 mol/kmol and even more preferably not higher than 30
mol/kmol. In the most preferred embodiment of the invention no
comonomer is introduced into the first polymerization stage and the
first homo- or copolymer of ethylene is a first homopolymer of
ethylene.
Second Polymerization Stage
In the second polymerization stage, a first polymer mixture
comprising the first homo- or copolymer of ethylene and a second
homopolymer of ethylene is formed. This is done by introducing the
particles of the first homo- or copolymer, containing active
catalyst dispersed therein, together with additional ethylene into
the second polymerization stage. Hydrogen is introduced for
controlling the molecular weight. This causes the second
homopolymer to form on the particles containing the first homo- or
copolymer.
The melt flow rate MFR.sub.2 of the first polymer mixture is from
10 to 250 g/10 min, preferably from 5 to 150 g/10 min, and most
preferably from 10 to 100 g/10 min. The MFR.sub.2 of the first
polymer mixture is higher than the MFR.sub.2 of the first homo- or
copolymer.
As it is well understood by the person skilled in the art, the melt
flow rate of the second homopolymer of ethylene cannot be measured
because the second homopolymer never exists in isolation. It is
also questionable if a mixing rule to calculate the melt flow rate
can be used, because a mixing rule fitted within the specific split
and MFR ranges does not exist. However, when producing the second
homopolymer in the second polymerization stage it is beneficial to
adjust the hydrogen to ethylene ratio within the fluid phase to
such values which would produce a homopolymer having a melt flow
rate of about from 100 to 2000 g/10 min, preferably about from 300
to 1000 g/10 min in a single-stage polymerization with the actual
catalyst.
The density of the first polymer mixture is typically higher than
965 kg/m.sup.3 and preferably from 968 to 978 kg/m.sup.3, more
preferably from 970 to 978 kg/m.sup.3. As it is understood by the
person skilled in the art the density of the first polymer mixture
cannot be controlled in the case both the first homopolymer and the
second homopolymer are ethylene homopolymers. Thereby the density
of the first polymer mixture is set by the melt indices of the
first and second homopolymers and the relative proportions of said
homopolymers.
Furthermore, the density of the second homopolymer cannot be
measured because the second homopolymer cannot be isolated from the
first polymer mixture. However, the density of the second
homopolymer can be estimated from the densities of the first
polymer mixture and the first homo- or copolymer by using the
equation: .rho..sub.b=w.sub.1.rho..sub.1+w.sub.2.rho..sub.2 (eq.
1)
Where .rho. is the density in kg/m.sup.3, w is the weight fraction
of the component in the mixture and subscripts b, 1 and 2 refer to
the overall mixture (=first polymer mixture), component 1 (=first
homo- or copolymer) and component 2 (=second homopolymer),
respectively.
The polymerization in the second polymerization stage is
advantageously conducted as a slurry polymerization at a
temperature of from 50 to 115.degree. C., preferably from 90 to
105.degree. C. The pressure is from 1 to 150 bar, preferably from
50 to 100 bar. Furthermore, it is preferred to conduct the
polymerization above the critical temperature and pressure of the
fluid mixture. Such operation is described in U.S. Pat. No.
5,391,654. In such operation the temperature is typically from 85
to 110.degree. C., preferably from 90 to 105.degree. C. and the
pressure is from 40 to 150 bar, preferably from 50 to 100 bar.
Hydrogen feed is adjusted to achieve a desired melt flow rate (or
molecular weight) of the first polymer mixture. Suitably, the
hydrogen feed is controlled to maintain constant hydrogen to
ethylene molar ratio in the reaction mixture. The actual ratio
depends on the catalyst as well as the type of the polymerization.
Good results have been obtained in slurry polymerization above the
critical temperature and pressure of the fluid mixture by
maintaining the ratio within the range of from 200 to 1000
mol/kmol, preferably from 300 to 800 mol/kmol.
Preferably, the second polymerization stage is conducted in a loop
reactor in a similar manner as was described for the first
polymerization stage except for what has been otherwise mentioned
above.
The first polymer mixture comprises from 2 to 20% by weight of the
first homo- or copolymer and from 80 to 98% by weight of the second
homopolymer. Preferably, the first polymer mixture comprises from 3
to 30% by weight of the first homo- or copolymer and from 70 to 97%
by weight of the second homopolymer. And most preferably, the first
polymer mixture comprises from 4 to 18% by weight of the first
homo- or copolymer and from 82 to 96% by weight of the second
homopolymer.
Third Polymerization Stage
In the third polymerization stage, a second polymer mixture
comprising the first polymer mixture and a third copolymer of
ethylene is formed. This is done by introducing the particles of
the first polymer mixture, containing active catalyst dispersed
therein, together with additional ethylene and comonomer into the
third polymerization stage. Hydrogen may be introduced for
controlling the molecular weight. This causes the third copolymer
to form on the particles containing the first polymer mixture.
The melt flow rate MFR.sub.21 of the second copolymer mixture is
from 1 to 50 g/10 min, preferably from 2 to 50 g/10 min, and most
preferably from 3 to 30 g/10 min. The melt flow rate of the second
copolymer mixture is lower than the melt index of the first polymer
mixture.
As explained above for the first polymer mixture, the MFR.sub.21 of
the third copolymer of ethylene cannot be measured because the
third copolymer cannot be isolated from the second polymer mixture.
However, the MFR.sub.21 of the third copolymer of ethylene can be
calculated by using the so called Hagstrom equation (Hagstrom, The
Polymer Processing Society, Europe/Africa Region Meeting,
Gothenburg, Sweden, Aug. 19-21, 1997).
.times. ##EQU00001##
As proposed by Hagstrom, a=10.4 and b=0.5 for MFR.sub.21. Further,
unless other experimental information is available,
MFR.sub.21/MFR.sub.2 for one polymer component can be taken as 30.
Furthermore, w is the weight fraction of the polymer component
having higher MFR. The first polymer mixture can thus be taken as
the component 1 and the third copolymer as the component 2. The
MFR.sub.21 of the third copolymer (MI.sub.2) can then be solved
from equation 1 when the MFR.sub.21 of the first polymer mixture
(MI.sub.1) and the second polymer mixture (MI.sub.b) are known.
The density of the third copolymer cannot be directly measured.
However, by using the standard mixing rule of equation 1 above it
can be calculated from the densities of the second polymer mixture
and the first polymer mixture. Then the subscripts b, 1 and 2 refer
to the overall mixture (=second polymer mixture), component 1
(=first polymer mixture) and component 2 (=third copolymer),
respectively.
Hydrogen feed is adjusted to achieve a desired melt flow rate (or
molecular weight) of the second polymer mixture. Suitably the
hydrogen feed is controlled to maintain constant hydrogen to
ethylene ratio in the reaction mixture. The actual ratio depends on
the catalyst as well as the type of the polymerization. Good
results have been obtained in gas phase polymerization by
maintaining the ratio within the range of from 2 to 100 mol/kmol,
preferably from 5 to 50 mol/kmol.
The comonomer is selected from ethylene and alpha-olefins
containing 4 to 10 carbon atoms and their mixtures. The content of
the comonomer is controlled to obtain the desired density of the
second polymer mixture. Typically the second polymer mixture has a
density of from 935 to 960 kg/m.sup.3, preferably from 940 to 955
kg/m.sup.3 and more preferably from 945 to 955 kg/m.sup.3.
The comonomer is typically introduced to maintain a constant
comonomer to ethylene ratio in the reaction mixture. The comonomer
to ethylene ratio that is needed to produce a polymer with the
desired density depends, among others, on the type of comonomer and
the type of catalyst. With 1-hexene as a comonomer good results
have been obtained in gas phase polymerization with a molar ratio
of 1-hexene to ethylene of from 0.1 to 200 mol/kmol, preferably
from 0.1 to 20 mol/kmol.
Preferably the third polymerization stage is conducted as a
fluidized bed gas phase polymerization. In a fluidized bed gas
phase reactor an olefin is polymerized in the presence of a
polymerization catalyst in an upwards moving gas stream. The
reactor typically contains a fluidized bed comprising the growing
polymer particles containing the active catalyst located above a
fluidization grid.
The polymer bed is fluidized with the help of the fluidization gas
comprising the olefin monomer, eventual comonomer(s), eventual
chain growth controllers or chain transfer agents, such as
hydrogen, and eventual inert gas. The fluidization gas is
introduced into an inlet chamber at the bottom of the reactor. To
make sure that the gas flow is uniformly distributed over the
cross-sectional surface area of the inlet chamber the inlet pipe
may be equipped with a flow dividing element as known in the art,
e.g. U.S. Pat. No. 4,933,149 and EP-A-684871. One or more of the
above-mentioned components may be continuously added into the
fluidization gas to compensate for losses caused, among other, by
reaction or product withdrawal.
From the inlet chamber the gas flow is passed upwards through a
fluidization grid into the fluidized bed. The purpose of the
fluidization grid is to divide the gas flow evenly through the
cross-sectional area of the bed. Sometimes the fluidization grid
may be arranged to establish a gas stream to sweep along the
reactor walls, as disclosed in WO-A-2005/087361. Other types of
fluidization grids are disclosed, among others, in U.S. Pat. No.
4,578,879, EP 600414 and EP-A-721798. An overview is given in
Geldart and Bayens: The Design of Distributors for Gas-fluidized
Beds, Powder Technology, Vol. 42, 1985.
The fluidization gas passes through the fluidized bed. The
superficial velocity of the fluidization gas must be higher than
minimum fluidization velocity of the particles contained in the
fluidized bed, as otherwise no fluidization would occur. On the
other hand, the velocity of the gas should be lower than the onset
velocity of pneumatic transport, as otherwise the whole bed would
be entrained with the fluidization gas. The minimum fluidization
velocity and the onset velocity of pneumatic transport can be
calculated when the particle characteristics are known by using
common engineering practise. An overview is given, among others in
Geldart: Gas Fluidization Technology, J. Wiley & Sons,
1986.
When the fluidization gas is contacted with the bed containing the
active catalyst the reactive components of the gas, such as
monomers and chain transfer agents, react in the presence of the
catalyst to produce the polymer product. At the same time the gas
is heated by the reaction heat.
The unreacted fluidization gas is removed from the top of the
reactor and cooled in a heat exchanger to remove the heat of
reaction. The gas is cooled to a temperature which is lower than
that of the bed to prevent the bed from heating because of the
reaction. It is possible to cool the gas to a temperature where a
part of it condenses. When the liquid droplets enter the reaction
zone they are vaporised. The vaporisation heat then contributes to
the removal of the reaction heat. This kind of operation is called
condensed mode and variations of it are disclosed, among others, in
WO-A-2007/025640, U.S. Pat. No. 4,543,399, EP-A-699213 and
WO-A-94/25495. It is also possible to add condensing agents into
the recycle gas stream, as disclosed in EP-A-696293. The condensing
agents are non-polymerizable components, such as n-pentane,
isopentane, n-butane or isobutane, which are at least partially
condensed in the cooler.
The gas is then compressed and recycled into the inlet chamber of
the reactor. Prior to the entry into the reactor fresh reactants
are introduced into the fluidization gas stream to compensate for
the losses caused by the reaction and product withdrawal. It is
generally known to analyze the composition of the fluidization gas
and introduce the gas components to keep the composition constant.
The actual composition is determined by the desired properties of
the product and the catalyst used in the polymerization.
The catalyst may be introduced into the reactor in various ways,
either continuously or intermittently. Among others, WO-A-01/05845
and EP-A-499759 disclose such methods. Where the gas phase reactor
is a part of a reactor cascade the catalyst is usually dispersed
within the polymer particles from the preceding polymerization
stage. The polymer particles may be introduced into the gas phase
reactor as disclosed in EP-A-1415999 and WO-A-00/26258.
The polymeric product may be withdrawn from the gas phase reactor
either continuously or intermittently. Combinations of these
methods may also be used. Continuous withdrawal is disclosed, among
others, in WO-A-00/29452. Intermittent withdrawal is disclosed,
among others, in U.S. Pat. No. 4,621,952, EP-A-188125, EP-A-250169
and EP-A-579426.
The top part of the gas phase reactor may include a so called
disengagement zone. In such a zone the diameter of the reactor is
increased to reduce the gas velocity and allow the particles that
are carried from the bed with the fluidization gas to settle back
to the bed.
The bed level may be observed by different techniques known in the
art. For instance, the pressure difference between the bottom of
the reactor and a specific height of the bed may be recorded over
the whole length of the reactor and the bed level may be calculated
based on the pressure difference values.
Such a calculation yields a time-averaged level. It is also
possible to use ultrasonic sensors or radioactive sensors. With
these methods instantaneous levels may be obtained, which of course
may then be averaged over time to obtain a time-averaged bed
level.
Also antistatic agent(s) may be introduced into the gas phase
reactor if needed. Suitable antistatic agents and methods to use
them are disclosed, among others, in U.S. Pat. No. 5,026,795, U.S.
Pat. No. 4,803,251, U.S. Pat. No. 4,532,311, U.S. Pat. No.
4,855,370 and EP-A-560035. They are usually polar compounds and
include, among others, water, ketones, aldehydes and alcohols.
The reactor may also include a mechanical agitator to further
facilitate mixing within the fluidized bed. An example of suitable
agitator design is given in EP-A-707513.
Typically the fluidized bed polymerization reactor is operated at a
temperature within the range of from 50 to 100.degree. C.,
preferably from 65 to 90.degree. C. The pressure is suitably from
10 to 40 bar, preferably from 15 to 30 bar.
The second ethylene polymer mixture comprises from 40 to 60% by
weight, preferably from 45 to 55% by weight of the first ethylene
polymer mixture, and from 40 to 60% by weight, preferably 45 to 55%
by weight of the third copolymer of ethylene.
The second polymer mixture preferably comprises from 45 to 60% by
weight, more preferably from 48 to 55% by weight of the first
polymer mixture, and from 40 to 55% by weight, more preferably from
45 to 52% by weight of the third copolymer of ethylene.
Extrusion
When the polymer has been removed from the polymerization reactor
it is subjected to process steps for removing residual hydrocarbons
from the polymer. Such processes are well known in the art and can
include pressure reduction steps, purging steps, stripping steps,
extraction steps and so on. Also combinations of different steps
are possible.
According to one preferred process a part of the hydrocarbons is
removed from the polymer powder by reducing the pressure. The
powder is then contacted with steam at a temperature of from 90 to
110.degree. C. for a period of from 10 minutes to 3 hours.
Thereafter the powder is purged with inert gas, such as nitrogen,
over a period of from 1 to 60 minutes at a temperature of from 20
to 80.degree. C.
According to another preferred process the polymer powder is
subjected to a pressure reduction as described above. Thereafter it
is purged with an inert gas, such as nitrogen, over a period of
from 20 minutes to 5 hours at a temperature of from 50 to
90.degree. C. The inert gas may contain from 0.0001 to 5%,
preferably from 0.001 to 1%, by weight of components for
deactivating the catalyst contained in the polymer, such as
steam.
The purging steps are preferably conducted continuously in a
settled moving bed. The polymer moves downwards as a plug flow and
the purge gas, which is introduced to the bottom of the bed, flows
upwards.
Suitable processes for removing hydrocarbons from polymer are
disclosed in WO-A-02/088194, EP-A-683176, EP-A-372239, EP-A-47077
and GB-A-1272778.
After the removal of residual hydrocarbons the polymer is
preferably mixed with additives as it is well known in the art.
Such additives include antioxidants, process stabilizers,
neutralizers, lubricating agents, nucleating agents, pigments and
so on.
The polymer particles are mixed with additives and extruded to
pellets as it is known in the art. Preferably a counter-rotating
twin screw extruder is used for the extrusion step. Such extruders
are manufactured, for instance, by Kobe and Japan Steel Works. A
suitable example of such extruders is disclosed in
EP-A-1600276.
Benefits of the Invention
The process of the invention is capable of producing multimodal
ethylene polymers having a combination of excellent processability
and good mechanical properties. Furthermore, the process of the
invention allows stable and smooth production of the multimodal
resin.
In more specific, the resulting polymers are useful for producing
pipes having improved processability and improved pressure
resistance. The resins furthermore have a reduced number and size
of white spots compared to similar prior art resins.
The polymer powder formed in the process of the present invention
has good handling properties. This allows a stable operation of the
process.
Description of Methods
Melt Flow Rate
Melt flow rate (MFR) was determined according to ISO 1133 at
190.degree. C. The load under which the measurement is conducted is
given as a subscript. Thus, the MFR under the load of 2.16 kg is
denoted as MFR.sub.2. The melt flow rate MFR.sub.21 is
correspondingly determined at 190.degree. C. under a load of 21.6
kg.
The melt index MFR is herein assumed to follow the following mixing
rule (equation 2):
.times. ##EQU00002##
Where a=10.4 and b=0.5, w is the weight fraction of the component
having the higher MFR in the mixture, MI is the melt index
MFR.sub.21 and subscripts b, 1 and 2 refer to the mixture,
component 1 and component 2, respectively.
Density
Density of the polymer was measured according to ISO 1183-1:2004
Method A on compression moulded specimen prepared according to EN
ISO 1872-2 (February 2007) and is given in kg/m.sup.3.
The density is herein assumed to follow the mixing rule (equation
1): .rho..sub.b=w.sub.1.rho..sub.1+w.sub.2.rho..sub.2 (eq. 1)
Where .rho. is the density in kg/m.sup.3, w is the weight fraction
of the component in the mixture and subscripts b, 1 and 2 refer to
the overall mixture, component 1 and component 2, respectively.
Flowability of the Polymer
The flowability of powder samples was measured by using a steel
funnel according to ASTM B417. The amount of the powder used in the
test was 50 grams. The powder was then allowed freely to discharge
from the funnel. The time when all the powder had discharged was
recorded and reported as flowability.
Activity of the Catalyst
The activity of the catalyst (in kg/(gh)) was determined as the
amount of polymer produced in the process (in kg/h) divided by the
amount of catalyst fed into the process (in g/h) per unit of the
residence time in the reactor (h).
Particle Size of the Catalyst
Particle size is measured by Coulter Counter LS 200 at room
temperature with n-heptane as medium.
EXAMPLES
Catalyst Preparation
Complex Preparation:
87 kg of toluene was added into the reactor. Then 45.5 kg Bomag A
(Butyloctyl magnesium) in heptane was also added in the reactor.
161 kg 99.8% 2-ethyl-1-hexanol was then introduced into the reactor
at a flow rate of 24-40 kg/h. The molar ratio between BOMAG-A and
2-ethyl-1-hexanol was 1:1.83.
Solid Catalyst Component Preparation:
330 kg silica (calcined silica, Sylopol.RTM. 2100, having median
particle size d.sub.50 of about 9 .mu.m) and pentane (0.12 kg/kg
carrier) were charged into a catalyst preparation reactor. Then
EADC (Ethylaluminium dichloride) (2.66 mol/kg silica) was added
into the reactor at a temperature below 40.degree. C. during two
hours and mixing was continued for one hour. The temperature during
mixing was 40-50.degree. C. Then Mg complex prepared as described
above was added (2.56 mol Mg/kg silica) at 50.degree. C. during two
hours and mixing was continued at 40-50.degree. C. for one hour.
0.84 kg pentane/kg silica was added into the reactor and the slurry
was stirred for 4 hours at the temperature of 40-50.degree. C.
Finally, TiCl.sub.4 (1.47 mol/kg silica) was added during at least
1 hour at 55.degree. C. to the reactor. The slurry was stirred at
50-60.degree. C. for five hours. The catalyst was then dried by
purging with nitrogen.
Molar composition of the ready catalyst is:
Al/Mg/Ti=1.5/1.4/0.8 (mol/kg silica).
Example 1
Polymers were synthesized by using a bench scale 5 L reactor and
following a three step process.
Into a stirred 5 liter reactor maintained at a temperature of
23.degree. C. were introduced 650 g propane diluent and
triethylaluminium (TEA) as a cocatalyst was introduced so that the
ratio of TEA/Ti was 15. The temperature was then raised to
40.degree. C. and 37.5 mg of the solid catalyst component prepared
as discussed above were added with 200 g of propane diluent,
together with 20 g ethylene. The temperature was then raised to
50.degree. C. after which ethylene was continuously introduced so
much that the pressure in the reactor was 22 barg. The
polymerization was continued until 35 g polymer were produced,
after which the temperature in the reactor was increased to
85.degree. C. and 53 L hydrogen was added to the reactor. Ethylene
was then fed into the reactor so that the pressure in the reactor
was 58.5 barg. The polymerization was continued until 232.5 g
polymer were produced. The polymerization was then interrupted by
venting the hydrocarbons from the reactor.
Then, the temperature of the reactor was adjusted to 80.degree. C.
and into the reactor were introduced 190 g propane diluent, 1.8 g
1-hexene as a comonomer and then a continuous C.sub.6/C.sub.2
feeding of 4 wt %, 3 mL hydrogen and ethylene so that the pressure
in the reactor was 20 barg. The polymerization was conducted until
232.5 g polymer were produced after which the hydrocarbons were
removed from the reactor and the powder was recovered and
dried.
The polymer was then extruded to pellets after addition of 2.5%
carbon black by using a twin-screw co-rotating extruder ZSK-18. The
data is shown in Table 1.
Examples 2 and 3
The procedure of Example 1 was repeated but the duration of the
first polymerization stage was adjusted so that down to 12.5 g
polymer was produced. The data is shown in Table 1.
Comparative Example 1
The procedure of Example 1 was repeated except that the first
polymerization stage was not present and temperature was increased
to 95.degree. C. right after the catalyst components were
introduced directly into the second polymerization stage. The data
is shown in Table 1.
TABLE-US-00001 TABLE 1 Example 1 2 3 CE1 Activity of the catalyst
5.3 5.4 6.0 5.9 [Kg PE/g cat/h] MFR.sub.5 0.08 0.09 0.11 0.15 [g/10
min] MFR.sub.21 4.6 5.34 5.0 7.1 [g/10 min] MFR.sub.5 0.08 0.1 0.14
0.2 [g/10 min] compounded MFR.sub.21 4.6 6.0 6.3 9.5 [g/10 min]
compounded FRR compounded 58 60 45 47 1-hexene content, 0.8 0.9 0.8
1.0 [wt-%] Density 951 952 951 952 [kg/m.sup.3] Split 7/46/47
5/47/48 2.5/48.5/49 0/51/49 1.sup.st/2.sup.nd/3.sup.rd
Examples 4 and 5
The procedure of Example 1 was repeated except that the solid
catalyst component was changed to a commercially available catalyst
Lynx 200 from BASF, and the pressure in the reactor in the first
polymerisation step was 20.3 barg, and the temperature of the
reactor in the second polymerisation step was adjusted to
95.degree. C., and the temperature of the reactor in the third
polymerisation step was 85.degree. C. and the continuous
C.sub.6/C.sub.2 feeding was of 7.5 wt %. The polymer was then
extruded to pellets after addition of 2.5% carbon black by using a
twin-screw co-rotating extruder ZSK-18. The data is shown in Table
2.
Comparative Example 2
The procedure of Comparative Example 1 was repeated except that the
solid catalyst component was changed to a commercially available
catalyst Lynx 200 from BASF and the temperature of the reactor in
the second polymerisation step was adjusted to 95.degree. C., and
the temperature of the reactor in the third polymerisation step was
85.degree. C. and the continuous C.sub.6/C.sub.2 feeding was of 7.5
wt % and 250 mL hydrogen were feed so that the pressure in the
reactor was 20.6 barg. The polymer was then extruded to pellets
after addition of 2.5% carbon black by using a twin-screw
co-rotating extruder ZSK-18. The data is shown in Table 2.
TABLE-US-00002 TABLE 2 Example 4 5 CE2 Activity of the catalyst 15
15 19 [Kg PE/g cat/h] MFR.sub.5 0.09 0.13 0.31 [g/10 min]
MFR.sub.21 5.0 7.7 13 [g/10 min] MFR.sub.5 0.07 0.07 0.31 [g/10
min] compounded MFR.sub.21 3.9 4.2 13 [g/10 min] compounded FRR
compounded 55 59 41 1-hexene content, 2.6 2.6 3.6 [wt-%] Density
(base resin) 948 947 948 [kg/m.sup.3] Density (compound) 959 959
961 [kg/m.sup.3] Split 9/46/45 4/49/47 0/51/49
1.sup.st/2.sup.nd/3.sup.rd
From the polymers of Examples 1 to 3 and Comparative Example 1,
mechanical properties were evaluated. They are shown in Table
3.
TABLE-US-00003 TABLE 3 Example 4 5 CE2 Tensile Modulus 1180 1180
1180 [MPa] Tensile stress at yield 27 27 26 [MPa] Tensile stress at
break 31 31 28 [MPa] Natural Draw Ratio 230 250 260 [%]
Examples 6 to 8
The procedure of Example 1 was repeated except that the first
polymerization step was conducted at temperatures 60.degree. C.,
40.degree. C. and 40.degree. C., respectively. The pressure in the
reactor in the first polymerisation step was 26.5 barg and 18 barg,
at 60.degree. C. and 40.degree. C. respectively. The data is shown
in table 4.
Comparative Example 3
The procedure of Comparative Example 1 was repeated except that
continuous C.sub.6/C.sub.2 feeding in the third polymerization step
was of 2.5 wt %. Also the relative amounts of the components were
changed by changing the polymerisation time in the first and second
step so that different amounts of polymers were produced. The data
is shown in Table 4.
The flowability of the polymer powders as described above was
measured. The data is also shown in Table 4.
TABLE-US-00004 TABLE 4 Example 6 7 8 CE3 Activity of the catalyst
5.6 4.4 4.2 5.8 [Kg PE/g cat/h] MFR.sub.5 0.08 0.06 0.08 0.16 [g/10
min] MFR.sub.21 3.6 2.7 5.2 8.2 [g/10 min] MFR.sub.5 0.06 0.07 0.08
0.2 [g/10 min] compounded MFR.sub.21 2.7 3.1 5.2 10 [g/10 min]
compounded FRR compounded 45 44 65 51 Split 10/45/45 9/46/45
8/53/39 0/55/45 1.sup.st/2.sup.nd/3.sup.rd Flowability 6 6 5 No
result [s]
No result indicates that the powder did not flow through the funnel
in the flowability test but jammed in the funnel.
Example 9
A loop reactor having a volume of 150 dm.sup.3 was operated at a
temperature of 50.degree. C. and a pressure of 60 bar. Into the
reactor were ethylene, propane diluent and hydrogen so that the
ethylene concentration in the fluid mixture was 4.4% by mole, the
hydrogen to ethylene ratio was 10 mol/kmol and the propane feed was
40 kg/h. Also 10 g/h of a solid polymerisation catalyst component
produced as described above was introduced into the reactor
together with triethylaluminium cocatalyst so that the molar ratio
of aluminium to titanium was 9 mol/mol. The ethylene homopolymer
withdrawn from the reactor had MFR.sub.21 of 0.4 g/10 min. The
production rate was 6 kg/h.
The slurry from the reactor was withdrawn intermittently and
directed into a loop reactor having a volume of 350 dm.sup.3 and
which was operated at 95.degree. C. temperature and 54 bar
pressure. Into the reactor was further added 100 kg/h of propane
and ethylene and hydrogen so that the ethylene content in the
reaction mixture was 2.8 mol-% and the molar ratio of hydrogen to
ethylene was 570 mol/kmol. The ethylene homopolymer withdrawn from
the reactor had a melt flow rate MFR.sub.2 of 35 g/10 min and the
production rate was 38 kg/h.
The slurry was withdrawn from the loop reactor intermittently by
using settling legs and directed to a flash vessel operated at a
temperature of 50.degree. C. and a pressure of 3 bar. From there
the polymer was directed to a gas phase reactor operated at a
pressure of 20 bar and a temperature of 85.degree. C. Additional
ethylene, 1-hexene comonomer, nitrogen as inert gas and hydrogen
were added so that the ethylene content in the reaction mixture was
12 mol-%, the ratio of hydrogen to ethylene was 10 mol/kmol and the
feed of 1-hexene to the reactor was 2 g/h. The polymer production
rate in the gas phase reactor was 40 kg/h and thus the total
polymer withdrawal rate from the gas phase reactor was 54 kg/h. The
polymer had a melt flow rate MFR.sub.5 of 0.12 g/10 min and a
density of 949 kg/m.sup.3. The production split (% 1.sup.st stage
component/% 2.sup.nd stage component/% 3.sup.rd stage component)
was 7/45/48.
The polymer was mixed with effective amounts of Ca-stearate,
Irganox B225 and HE0880 (Carbon Black masterbatch, corresponding to
2.25% by weight of carbon black). Then it was compounded and
extruded under nitrogen atmosphere to pellets by using a CIMP90
extruder so that the SEI was 233 kWh/ton and the melt temperature
273.degree. C. under nitrogen atmosphere.
Example 10
The procedure of Example 9 was followed except that the operation
conditions in the loop reactor and the gas phase reactor were
modified as shown in Table 5. The powder was not compounded but the
material was recovered in powder form.
TABLE-US-00005 TABLE 5 Example 9 10 1.sup.st loop C.sub.2, [mol-%]
4.4 4.3 1.sup.st loop H.sub.2/C.sub.2, [mol/kmol] 10 18 1.sup.st
loop MFR.sub.21, [g/10 min] 0.4 N/D.sup.3) 2.sup.nd loop C.sub.2,
[mol-%] 2.8 3.3 2.sup.nd loop H.sub.2/C.sub.2, [mol/kmol] 570 500
2.sup.nd loop MFR.sub.2, [g/10 min] 35 37 GPR, C.sub.2, [mol-%] 12
15 GPR H.sub.2/C.sub.2, [mol/kmol] 10 13 GPR C.sub.6/C.sub.2,
[mol/kmol] N/D.sup.2) N/D.sup.2) GPR MFR.sub.5, [g/10 min]
0.12.sup.1) 0.16.sup.1) GPR density, [kg/m.sup.3] 949.sup.1) .sup.
951.sup.1) Compound MFR.sub.5, [g/10 min] 0.16 N/D.sup.3) Compound
MFR.sub.21, [g/10 min] 4.3 3.9.sup.1) Compound density,
[kg/m.sup.3] 963 N/D.sup.3) Notes: .sup.1)measured from powder
.sup.2)C.sub.6 peak not detected by online GC .sup.3)Not
analyzed
* * * * *